U.S. patent number 6,252,126 [Application Number 09/218,840] was granted by the patent office on 2001-06-26 for method for producing ethylbenzene.
Invention is credited to David Netzer.
United States Patent |
6,252,126 |
Netzer |
June 26, 2001 |
Method for producing ethylbenzene
Abstract
An ethylbenzene production system comprises a reactor vessel, a
vapor phase ethylene feed stream, a benzene feed stream entering
the reactor vessel, and a product stream containing ethylbenzene
exiting the reactor vessel. The reactor vessel has an ethylation
section and a benzene stripping section, whereby fluid
communication via integrated vapor and liquid traffic is maintained
between the ethylation section and stripping section. The vapor
phase ethylene feed stream contains 3 to 50 mol % ethylene and at
least 20 mol % methane entering the reactor vessel.
Inventors: |
Netzer; David (Los Angeles,
CA) |
Family
ID: |
27376379 |
Appl.
No.: |
09/218,840 |
Filed: |
December 22, 1998 |
Related U.S. Patent Documents
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Application
Number |
Filing Date |
Patent Number |
Issue Date |
|
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175643 |
Oct 20, 1998 |
5977423 |
|
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|
Current U.S.
Class: |
585/446; 562/607;
585/323; 422/652; 422/623; 585/467 |
Current CPC
Class: |
C07C
15/073 (20130101); C07C 7/00 (20130101); C07C
7/09 (20130101); B01D 3/009 (20130101); C07C
7/09 (20130101); C07C 15/04 (20130101); C07C
7/00 (20130101); C07C 15/04 (20130101); Y02P
20/10 (20151101); Y02P 20/127 (20151101) |
Current International
Class: |
B01D
3/00 (20060101); C07C 7/09 (20060101); C07C
15/073 (20060101); C07C 7/00 (20060101); C07C
15/00 (20060101); B01J 010/00 (); C07C
002/64 () |
Field of
Search: |
;260/669R ;422/189
;562/607 ;585/323,467,446 |
References Cited
[Referenced By]
U.S. Patent Documents
Foreign Patent Documents
Primary Examiner: Knode; Marian C.
Assistant Examiner: Varcoe; Frederick
Attorney, Agent or Firm: Christie, Parker & Hale,
LLP
Parent Case Text
CROSS-REFERENCE TO RELATED APPLICATIONS
This application is a continuation-in-part of U.S. application Ser.
No. 09/175,643, filed on Oct. 20, 1998, now U.S. Pat. No.
5,977,423, which claims priority of U.S. Provisional Application
No. 60/089,968, filed on Jun. 19, 1998, the disclosures of which
are incorporated herein by reference.
Claims
What is claimed is:
1. A method for producing ethylbenzene comprising:
operating an isothermal ethylation section of a reactor vessel in a
steady state mode, the isothermal ethylation section comprising
catalytic media, the reactor vessel further comprising a benzene
stripping section, where the stoichiometric excess, unreacted
benzene is thermally stripped from the products, whereby integrated
vapor and liquid traffic is maintained between the ethylation
section and the benzene stripping section;
providing a vapor phase dilute ethylene feed stream entering the
reactor vessel comprising hydrogen, methane, and ethylene in a
concentration of from about 3 to 50 mol % based on the total
concentration of the ethylene feed stream;
providing a benzene feed stream entering the reactor vessel
comprising benzene and at least one non-aromatic compound wherein
the concentration of benzene in the benzene feed stream is from
about 75% to about 100% by weight, based on the total weight of
benzene and non-aromatic compounds, and wherein the amount of
benzene in the ethylation section results in stoichiometric excess
of benzene relative to ethylene; and
reacting the ethylene with the benzene to produce a product stream
containing ethylbenzene and an overhead gas comprising hydrogen,
methane, unreacted ethylene, ethylbenzene and benzene exiting the
ethylation section of the reactor vessel, wherein the ethylation
reaction occurs at a temperature at least 10.degree. C. below the
boiling point of benzene at the pressure at which the reaction is
maintained;
wherein the isothermal ethylation section comprises a tubular
structure and maintains a substantially constant reaction
temperature in the catalytic media during the reaction of ethylene
with benzene by simultaneous removal of heat from the catalytic
reaction media via heat transfer across the surface of the tubular
structure.
2. A method according to claim 1, wherein the concentration of
benzene in the benzene feed stream is from about 92% to about 100%
by weight, based on the total weight of benzene and non-aromatic
compounds.
3. A method as claimed in claim 1 wherein the stoichiometric excess
of benzene in the benzene feed stream to ethylene in the ethylene
feed stream is less than 50%.
4. A method as claimed in claim 1 wherein the stoichiometric excess
of benzene in the benzene feed stream to ethylene in the ethylene
feed stream is less than 15%.
5. A method as claimed in claim 1 wherein the stoichiometric excess
of benzene in the benzene feed stream to ethylene in the ethylene
feed stream is less than 5%.
6. A method according to claim 1, further comprising providing a
purge reactor and feeding an impure benzene mixture produced as a
residual product in the ethylation section to the purge reactor,
wherein the impure benzene mixture comprises a higher concentration
of the at least one aromatic compound than the benzene feed stream
to the reactor vessel, thereby producing a second residual product
having a benzene concentration of below 30% by weight based on the
total weight of benzene and non-aromatics present in the
product.
7. A method as claimed in claim 1, wherein the temperature of the
ethylation section ranges from about 130.degree. C. to about
220.degree. C.
8. A method as claimed in claim 1, wherein the temperature of the
ethylation reaction ranges from about 155.degree. C. to about
195.degree. C.
9. A method as claimed in claim 1, wherein the operating pressure
inside the ethylation section ranges from about 7 to about 33
kg/cm2-g.
10. A method as claimed in claim 1, wherein the operating pressure
inside the ethylation section ranges from about 20 to about 26
kg/cm2-g.
11. A method as claimed in claim 1, wherein in the ethylation
section, steam is generated in steam tubes.
12. A method as claimed in claim 1, wherein the method produces an
ethylbenzene product containing less than 150 PPM by weight of
cumene.
13. A method according to claim 1, wherein the ethylene feed stream
comprises 3 to 50 mol % ethylene and at least 20 mol % methane.
14. A method according to claim 1, wherein the catalytic media is
contained within the tubular structure, whereby heat is transferred
through the surface of the tubular structure, thereby generating
steam outside the tubular structure.
15. A method according to claim 1, wherein the catalytic media is
contained outside the tubular structure, whereby heat is
transferred through the surface of the tubular structure, thereby
generating steam inside the tubular structure.
16. A method according to claim 1, wherein the benzene in the
ethylation section of the reactor vessel is in a mixed vapor/liquid
phase.
17. A method according to claim 1, wherein the overhead gas
proceeds to a rectifying section of the reactor vessel,
ethylbenzene is recovered by reflux wash with benzene, and residual
vent gas comprising hydrogen, methane, unreacted ethylene and
benzene proceeds to benzene recovery by condensation.
18. A method for producing ethylbenzene comprising:
operating an ethylation section of a reactor vessel in a steady
state mode, wherein the reactor vessel comprises a mixed phase
transalkylation section and a benzene stripping section, where the
stoichiometric excess, unreacted benzene is thermally stripped from
the ethylation products, whereby fluid communication via integrated
vapor and liquid traffic is maintained between the ethylation
section and the transalkylation section and between the
transalkylation section and the benzene stripping section;
providing a vapor phase ethylene feed stream containing ethylene
entering the reactor vessel;
providing a benzene feed stream containing benzene entering the
reactor vessel, wherein the amount of benzene in the ethylation
section results in stoichiometric excess of benzene relative to
ethylene; and
reacting the ethylene with the benzene to produce a product stream
containing ethylbenzene and an overhead gas comprising hydrogen,
methane, unreacted ethylene, ethylbenzene and benzene exiting the
ethylation section of the reactor vessel, wherein the reaction
occurs at a temperature at least 10.degree. C. below the boiling
temperature of benzene at the pressure at which the reaction is
maintained.
19. A method according to claim 18, wherein the benzene in the
ethylation section of the reactor vessel is in a mixed vapor/liquid
phase.
20. A method according to claim 18, wherein the overhead gas
proceeds to a rectifying section of the reactor vessel,
ethylbenzene is recovered by reflux wash with benzene, and residual
vent gas comprising hydrogen, methane, unreacted ethylene and
benzene proceeds to benzene recovery by condensation.
21. A method according to claim 20, further comprising introducing
a source of at least one non-aromatic C.sub.5 to C.sub.7
hydrocarbon to the benzene feed stream or fractionation system to
depress the solid formation temperature of the overhead gas to
below 5.5.degree. C.
22. A method according to claim 20, further comprising introducing
a source of at least one non-aromatic C.sub.5 to C.sub.7
hydrocarbon to the benzene feed stream or fractionation system to
depress the solid formation temperature of the vent gas to below
about 5.5.degree. C.
23. A method according to claim 18, wherein the ethylation section
is an isothermal ethylation section comprising catalytic media,
wherein the isothermal section comprises a tubular structure and
maintains a substantially constant reaction temperature in the
catalytic media during the reaction of ethylene with benzene by
simultaneous removal of heat from the catalytic reaction media via
heat transfer across the surface of the tubular structure.
24. A method according to claim 23, wherein the catalytic media is
contained within the tubular structure, whereby heat is transferred
through the surface of the tubular structure, thereby generating
steam outside the tubular structure.
25. A method according to claim 23, wherein the catalytic media is
contained outside the tubular structure, whereby heat is
transferred through the surface of the tubular structure, thereby
generating steam inside the tubular structure.
26. A method according to claim 18, further comprising providing a
purge reactor and feeding an impure benzene mixture produced as a
residual product in the ethylation section to the purge reactor,
wherein the impure benzene mixture comprises a higher concentration
of the at least one aromatic compound than the benzene feed stream
to the reactor vessel, thereby producing a second residual product
having a benzene concentration of below 30% by weight based on the
total weight of benzene and non-aromatics present in the product.
Description
FIELD OF THE INVENTION
The present invention is directed to an improved process for
producing ethylbenzene.
PRIOR ART AND BACKGROUND OF THE INVENTION
Ethylbenzene, C.sub.8 H.sub.10, is a key raw material in the
production of styrene and is produced by the ethylation reaction of
ethylene, C.sub.2 H.sub.4, and benzene C.sub.6 H.sub.6 in a
catalytic environment. Old ethylbenzene production plants,
typically built before 1980, used AlCl.sub.3 or BF.sub.3 as acidic
catalysts. The newer plants in general have been switching to
zeolite-based acidic catalysts. The typical purity of the benzene
feed, known as nitration grade benzene, is 99.9 wt %. The typical
purity of the ethylene feed would exceed 99.9 mol %.
A significant source of crude benzene is pyrolysis gasoline
(C.sub.5 to C.sub.9), which typically contains 55-75 wt %
aromatics. Pyrolysis gasoline, produced in naphtha based or heavy
liquid based olefin plants, contains 35-55 wt % benzene. About 35%
of world's benzene production capacity originates from pyrolysis
gasoline. Typically, after pyrolysis gasoline is hydrotreated for
saturation of olefins and di-olefins, the pyrolysis gasoline (free
of olefins and sulfur compounds) is exported to battery limits for
aromatics extraction process. Pure benzene, 99.9 wt %, along with
toluene and xylene, is a typical product of aromatic
extraction.
Impure benzene, 94-98 wt %, which is a 75-83.degree. C. atmospheric
cut, can be recovered from hydrotreated pyrolysis gasoline by a
simple fractionation process, as described in U.S. Pat. No.
5,880,320, the disclosure of which is incorporated herein by
reference.
Three types of ethylation reactor systems are used for producing
ethylbenzene, namely, vapor phase reactor systems, liquid phase
reactor systems, and mixed phase reactor systems. In vapor-phase
reactor systems, the ethylation reaction of benzene and ethylene is
carried out at about 380-420.degree. C. and a pressure of 9-15
kg/cm.sup.2 -g. In most cases, these systems use ethylene feed in
pure form as produced in conventional olefin plants. Dilute
ethylene streams, about 10-15 vol %, as produced in fluid catalytic
cracking (FCC) in petroleum refining, are converted to ethylbenzene
using vapor phase reaction. One known facility was designed by
Raytheon Engineers & Constructors and is operated by Shell
Chemicals at UK. Similar facilities for FCC off-gases were built in
China by Sinopec.
Vapor phase reactor systems comprise multiple fixed beds of zeolite
catalyst. Ethylene exothermally reacts with benzene to form
ethylbenzene, although undesirable chain and side reactions also
occur. About 15% of the ethylbenzene formed further reacts with
ethylene to form di-ethylbenzene isomers (DEB), tri-ethylbenzene
isomers (TEB) and heavier aromatic products. All these chain
reaction products are commonly referred as polyethylated benzenes
(PEBs). In addition to the ethylation reactions (at times referred
to in the industry as alkylation reactions), the formation of
xylene isomers as trace products occurs by side reactions. This
xylene formation in vapor phase can yield an ethylbenzene product
with about 0.05-0.20 wt % of xylenes. The xylenes show up as an
impurity in the subsequent styrene product, and are generally
considered undesirable.
Additionally, traces of propylene may enter the system with the
ethylene feed or are formed by catalytic cracking of non-aromatic
impurities that may enter with the benzene feed. The presence of
propylene results in the formation of isopropyl benzene, commonly
known as cumene, which is very undesirable in the ethylbenzene at
concentrations above 150 PPM. The cracking of non-aromatic
impurities is accelerated by increasing the ethylation reaction
temperature, and thus substantial cracking of non-aromatic
impurities to propylene occurs if the ethylation or transalkylation
reaction is at temperatures of above 300.degree. C. and in presence
of acidic catalyst. This may result in an unacceptable level of
cumene in the ethylbenzene product.
In order to minimize the formation of PEBs, a stoichiometric excess
of benzene, about 400-900% per pass, is applied, depending on
process optimization. The effluent from the ethylation reactor
contains 70-85 wt % of unreacted benzene, 12-20 wt % of
ethylbenzene product and about 3-4 wt % of PEBs. The PEBs are
converted back to ethylbenzene to avoid a yield loss.
The effluent of the ethylation reactor can undergo ethylbenzene
product recovery by several multiple fractionation stages. Benzene
can be recovered in a benzene recovery column by stripping and can
be recycled to the ethylation reactor. Ethylbenzene product can be
recovered in an ethylbenzene recovery column. DEB and TEB can be
separated from heavier aromatics in a PEB column. The heavy
aromatics can be diverted to the fuel oil system.
The DEB and TEB mixture proceeds to a transalkylation reactor
system where stoichiometric excess (250-300%) of benzene reacts
with DEB and TEB in vapor phase at about 420-450.degree. C. About
60-70% of the PEB is converted to ethylbenzene per pass. The
effluent product of transalkylation reactor consists of
ethylbenzene, un-reacted benzene and unconverted PEBs. This
transalkylated stream undergoes stabilization for light ends
removal and is recycled to fractionation in the benzene column. The
ultimate conversion of DEB and TEB to ethylbenzene is essentially
100%.
The boiling point of the xylene isomer trace products is very close
to that of the ethylbenzene, and thus no practical separation is
possible. The ethylbenzene product typically contains 500-2,000 PPM
by weight of xylene isomers, as well as 1000-2,000 PPM by weight of
benzene.
In recent years the trend in industry has been to shift away from
vapor phase reactors to liquid phase reactors. Liquid phase
reactors operate about 260-270.degree. C., which is under the
critical temperature of benzene, 290.degree. C. One advantage of
the liquid phase reactor is the very low formation of xylenes and
oligomers. The rate of the ethylation reaction is lower compared
with the vapor phase, but the lower design temperature of the
liquid phase reaction usually economically compensates for the
negatives associated with the higher catalyst volume. The
stoichiometric excess of benzene in liquid phase systems is
150-400%, compared with 400-800% in vapor phase. However, due to
the kinetics of the lower ethylation temperatures, resulting from
the liquid phase catalyst, the rate of the chain reactions forming
PEBs is considerably lower; namely, about 5-8% of the ethylbenzene
is converted to PEBs in liquid phase reactions versus the 15-20%
converted in vapor phase reactions. Transalkylation reaction, where
polyethylated benzene reacts with benzene to form ethylbenzene, can
occur in a liquid phase or vapor phase system. The liquid phase
reaction temperature would be 230-270.degree. C. The fractionation
sequences and product recovery methods for liquid phase reaction
systems are similar to those used in connection with vapor phase
reactor systems.
In recent years, technology has been developed for the production
of ethylbenzene from dilute ethylene streams by a mixed phase
reactor. The demonstrated dilute ethylene stream sources are from
petroleum refineries, fluid catalytic cracking operation (FCC). ABB
Lummus Global and CDTech have developed a mixed phase process.
Aside from ethylation reactors, the sequence of the ethylbenzene
product recovery and transalkylation is similar to the conventional
liquid phase reactor systems.
A potentially alternate source of dilute ethylene is described in
U.S. Pat. No. 5,880,320. The dilute ethylene stream is extracted
from the demethanizer section of the ethylene plant at about 22-30
kg/cm.sup.2 -g. Dilute gas from ethylene plants may contain 7-25
mol % ethylene, and the bulk of the balance is methane and
hydrogen. The propylene content is controlled at the ethylene
source to remain below 20 PPM by volume.
The use of a liquid phase reaction system for dilute ethylene
streams is not possible. Due to the high methane and hydrogen
content in the ethylene stream, the bubble point temperature of the
combined mixture of dilute ethylene and benzene is very low, lower
than the activity temperature of the ethylation catalyst, and
actually below the freezing point of benzene.
The reaction temperature of the mixed phase ethylation reactor is
under the dew point of the dilute ethylene benzene mixture, but
well above the bubble point. The diluents of the ethylene feed
comprise hydrogen, methane and small amounts of ethane, and CO
remains essentially in the vapor phase. The benzene in the reactor
is split between vapor phase and liquid phase, and the ethylbenzene
and PEB reaction products remain essentially in liquid phase.
In the alkylation and transalkylation of aromatic hydrocarbons,
zeolite catalysts have been shown to be an adequate substitute for
acidic catalysts, such as aluminum chloride (AlCl.sub.3), boron
trifluoride (BF.sub.3), liquid and solid phosphoric acid, sulfuric
acid and the like. For example, U.S. Pat. No. 2,904,607 shows
alkylation of aromatics in the presence of a crystalline
aluminosilicate having a uniform pore opening of 6 to 15
angstroms.
U.S. Pat. No. 3,641,177 describes an alkylation process wherein the
catalyst has undergone a series of ammonium exchange, calcination
and steam treatments. This catalyst would currently be described as
an "ultrastable" or "steam-stabilized" zeolite Y catalyst.
U.S. Pat. Nos. 3,751,504 and 3,751,506 show transalkylation and
alkylation over ZSM-5 type catalysts. Use of other medium-pore to
large-pore zeolites are taught in U.S. Pat. No. 4,016,245 (ZSM-35),
U.S. Pat. No. 4,046,859 (ZSM-21), U.S. Pat. No. 4,070,407 (ZSM-35
and ZSM-38), U.S. Pat. No. 4,076,842 (ZSM-23) U.S. Pat. No.
4,575,605 (ZSM-23), U.S. Pat. No. 4,291,185 (ZSM-12), U.S. Pat. No.
4,387,259 (ZSM-12), and U.S. Pat. No. 4,393,262 (ZSM-12) and
European Patent Application Nos. 7,126 (zeolite omega) and 30,084
(ZSM-4, zeolite beta, ZSM-20, zeolite L).
Liquid phase alkylation is specifically taught using zeolite beta
in U.S. Pat. No. 4,891,458 and European Patent Application Nos.
0432814 and 0629549. Novel dealuminized mordenites are described
for these types of reactions in U.S. Pat. Nos. 5,015,797 and
4,891,448.
More recently it has been disclosed that MCM-22 and its structural
analogues have utility in these alkylation/transalkylation
reactions. U.S. Pat. No. 4,992,606 (MCM-22), U.S. Pat. No.
5,258,565 (MCM-36), U.S. Pat. No. 5,371,310 (MCM-49), U.S. Pat. No.
5,453,554 (MCM-56), and U.S. Pat. No. 5,149,894 (SSZ-25).
Additionally Mg APSO-31 is described as an attractive catalyst for
cumene manufacture in U.S. Pat. No. 5,434,326.
U.S. Pat. No. 5,176,883 describes an integrated ethylation
fractionation in general without diluants for the ethylene feed.
U.S. Pat. No. 5,043,506 describes the addition of n-C.sub.5,
n-C.sub.6, and i-C.sub.6 as a means for fractionation control in
alkylation systems.
SUMMARY OF THE INVENTION
In one embodiment, the present invention is directed to an
ethylbenzene production system comprising a reactor vessel, a vapor
phase ethylene feed stream and a benzene feed stream entering the
reactor vessel, and a product stream containing ethylbenzene
exiting the reactor vessel. The reactor vessel has an ethylation
section and a benzene stripping section, whereby integrated vapor
and liquid traffic is maintained between the ethylation section and
stripping section. Preferably the reactor vessel is a single unit,
but can alternatively comprise a plurality of integrated units, so
long as integrated vapor and liquid traffic is maintained between
the integrated units.
Dilute ethylene streams at a typical concentration of 7-25 mol %
and less than 20 PPM of propylene will react with benzene at
temperatures of 155-195.degree. C. and pressures of 22-30 kg/cm2-g
to form ethylbenzene and small amount of PEBs. The benzene stripper
generates an internal benzene traffic in the ethylation section of
about 300-400% stoichiometric excess, but additional vapor traffic
is generated by the heat of reaction. The external stoichiometric
excess of benzene is on the order of 3-15% depending on purge rate,
purge recovery, benzene losses to vent gas and formation of heavy
aromatic product. The stripper's bottom product is essentially free
of benzene and suitable for ethylbenzene fractionation.
Preferably the reactor vessel further comprises a rectifying
section and a transalkylation section. Alternatively, the system
can comprises a transalkylation section outside of the reactor
vessel.
In a particularly preferred embodiment, the invention is directed
to an ethylbenzene production system comprising a reactor vessel, a
vapor phase dilute ethylene feed stream and a benzene feed stream
entering the reactor vessel, and a product stream containing
ethylbenzene exiting the reactor vessel. The reactor vessel has an
ethylation section and a benzene stripping section, whereby
integrated vapor and liquid traffic is maintained between the
ethylation section and stripping section. The vapor phase dilute
ethylene feed stream comprises ethylene in a concentration of from
about 3 to 50 mol % based on the total concentration of the
ethylene feed stream. The benzene feed stream comprises benzene and
at least one non-aromatic compound, wherein the concentration of
benzene in the benzene feed stream is from about 75% to about 100%
by weight, based on the total weight of benzene and non-aromatic
compounds.
The methods of the invention are particularly useful for the
utilization of impure benzene, typically 94-98 wt %, with a balance
of cyclohexane and other non-aromatics. The source of this impure
benzene would be pyrolysis gasoline, after hydrogenation and
fractionation. Because the temperatures of the ethylation and
transalkylation reactions are below 300.degree. C., no significant
cracking of non-aromatic occurs. Production of xylene is very
minimal if any.
Conventional zeolitic and nonzeolitic catalysts, with formulations
in the public domain, can be used. These catalysts have been
traditionally used for cumene manufacturing in a temperature range
of 150-180.degree. C. In the cumene reaction, propylene reacts with
benzene to form isopropyl benzene; however, impurities of ethylene
are also known to react with benzene to form ethylbenzene. The
non-aromatic impurities are allowed to build in the ethylation
loop, prior to purging to a purge ethylation reactor. The economics
of the assumed purge reactor would largely depend on the
concentration of non-aromatic impurities in the benzene feed.
In another embodiment, the invention is directed to a method for
enhancing the recovery of benzene from an impure benzene feed.
Cyclohexane is included in the impure benzene feed. The temperature
of the impure benzene feed is then reduced below the freezing point
of benzene. Preferably the temperature of the impure benzene feed
is reduced to a temperature ranging from about -6.degree. C. to
about 4.degree. C., more preferably from about -5.degree. C. to
about 0.degree. C. Preferably the weight ratio of non-aromatics to
benzene in the impure benzene feed ranges from about 0.25 to about
0.7.
DESCRIPTION OF THE DRAWINGS
FIG. 1 illustrates one embodiment of the invention where the
transalkylation catalyst beds are an integrated section of the
overall ethylation reactor vessel.
FIG. 2 illustrates an alternative embodiment of the invention where
the transalkylation reactor is a liquid phase reactor and separated
from the main reactor vessel.
DETAILED DESCRIPTION OF THE INVENTION
One embodiment of an ethylbenzene production system according to
the invention is depicted in FIG. 1. The ethylbenzene production
system comprises a reactor vessel having several sections, namely
an ethylation section, 10B, a transalkylation section, 10C, a
rectifying section, 10A, and a benzene stripping section 10D. The
reactor vessel, although depicted as a single vessel in FIG. 1, can
be in the form of several integrated vessels, so long as the
integrated vapor and liquid traffic is maintained between the
ethylation and the stripping section.
The ethylation section, a fixed bed catalytic ethylation section,
where vapor phase ethylene and mixed phase benzene feed streams
react to form ethylbenzene and PEB, is an isothermal reactor.
Because of the diluent effect of the methane and hydrogen in the
ethylene feed, the ethylation reaction is carried out at a
temperature that is at least 10.degree. C. lower than the normal
boiling temperature of benzene at a given ethylation pressure. The
invention is not intentionally directed to the use of catalytic
distillation in the reactor. Cocurrent, mixed current and counter
current flows would be formed, however, the thermodynamic effect on
ethylene conversion by simultaneous separation of ethylbenzene
product is insignificant.
The heat of reaction and reboiler heat input to the system are
recovered as 3-8 kg/cm.sup.2 -g steam, to be generated in the tubes
when the catalyst is placed in the shell. However, when the
catalyst is placed in the tubes, the steam is generated in the
shell. Additional steam at 1.5-2.0 kg/cm2-g would be generated at
the overhead condenser. The dilute ethylene stream (containing
methane and hydrogen) is introduced at the bottom of the catalytic
ethylation section. The catalyst formulation is available at the
public domain from cumene manufacturing technology. AlCl.sub.3
catalysts, which are known to be active for ethylation reactions at
about 150.degree. C., could also be considered as a viable option
for this system. Hydrogen, methane, vapor phase benzene and
cyclohexane pass to the rectifying section, and ethylbenzene, PEB
products, liquid phase benzene and heavy aromatics pass to the
transalkylation section.
Fixed beds of catalyst will serve as transalkylator using a vapor
liquid mixture of benzene and impurities such as cyclohexane,
however no diluents of the ethylene feed the ethylbenzene and PEB
will be essentially in liquid phase. The catalyst formulation for
the transalkylation can be identical to the one used for the
ethylation section. The heat effect of this reaction is nearly
zero, and the operating temperature range would be 220-250.degree.
C. depending on the pressure. The stoichiometric excess of benzene
in the transalkylation section is over 1000%, thus over 50 percent
conversion of PEB to ethylbenzene per pass occurs for the end of
run.
The remaining PEBs (after transalkylation), along with the
ethylbenzene, benzene, and heavy aromatics proceed to the benzene
stripping section. The stripping section is at the bottom of the
reaction vessel and is the section where benzene stripping occurs.
About 25 actual trays (15 theoretical) or equivalent packing can be
used. Stripping duty is provided by a fired heater (or hot oil)
providing thermal duty at about 295-325.degree. C., depending on
the pressure. The unreacted benzene along with cyclohexane is
driven to the catalytic section, creating a localized excess of
benzene, which improves reaction equilibrium to minimize PEB
formation. The stripping heat input also increases the ethylation
reaction temperature, thus improving the ethylation rate of
reaction and minimizing the amount of catalyst required.
The upper section of the reactor vessel acts as a rectifier where
reflux of benzene washes down the ethylbenzene vapors for full
product recovery. Vent gas, depleted of ethylene, proceeds to
residual benzene recovery by refrigeration.
In a preferred embodiment, a purge reactor similar to the one of
the main reactor vessel described above, but without
transalkylation section, could be included in the system. A purge
stream from the ethylation loop with 60 to 85 wt % benzene (73% in
the demonstrated case) reacts in a mixed phase with dilute
ethylene. Because of the low benzene to ethylene ratio, the
conversion of ethylbenzene to PEBs may reach 50% or more. The
bottom of the stripping section consists of ethylbenzene 35-65 wt %
and the balance is PEBs along with traces of benzene. This stream
is routed to the feed of the ethylbenzene column as shown in FIG.
1.
The off gas from the purge reactor is chilled for benzene and
cyclohexane condensation. The off-gas, rich in unconverted
ethylene, proceeds to the ethylene feed stream of the main
ethylation reactor. The residual, non-converted liquid resulting
from the ethylation of the purge contains approximately 15-20 wt %
benzene and the balance non-aromatics, principally cyclohexane.
This liquid is disposed to the pyrolysis gasoline export, or to a
crude cyclohexane facility.
In an alternative embodiment, as shown in FIG. 2, the
transalkylation section is contained outside of the reaction
vessel. This alternative design can be used if there is a concern
of catalyst plugging and deactivation by heavy aromatics, or if the
reaction is being run at a pressure below 20 kg/cm.sup.2 g. The
resulting operating temperature of below 220.degree. C. would
deactivate the catalyst.
When an impure benzene feed is used (for example, a feed
originating from pyrolysis gasoline fractionation) the cyclohexane
concentration builds in the upper rectifying section. The freezing
points of pure benzene and cyclohexane are +5.5.degree. C. and
+6.degree. C., respectively. The eutectic effect of cyclohexane
buildup results in depression of the freezing temperature of the
mixture to approximately -17.degree. C. to -10.degree. C.
(depending on the ratio of benzene to cyclohexane). Thus, the
benzene mixture from the rectifying section can be cooled to a
temperature of -5.degree. C. or lower. The lower temperature
permits a greater amount of benzene recovery from the vent gas.
Thus benzene recovery by refrigeration of the vent gas is a
feasible approach, and the more conventional vent gas scrubber
using PEB liquid can be avoided. For pure benzene feed, a
conventional vent scrubber is required, unless cyclohexane is added
to the overhead.
In U.S. patent application Ser. No. 08/957,252, the usage of impure
benzene is proposed in conjunction with hydrotreating and
fractionation of pyrolysis gasoline in an adjacent olefin plant.
Typically, this impure benzene resulting from pyrolysis gasoline
includes about 2-6 wt % cyclohexane, depending on naphtha feed
analysis and cracking severity in the olefin plant. The
non-aromatic impurities also include traces of other C.sub.6 and
C.sub.7 's. These impurities would be allowed to build to a weight
ratio of 0.3-0.70 to the benzene in the reflux drum. The eutectic
effect of the impurities will allow the chilling the vent gas to
-2.degree. C. to -10.degree. C., becoming an economical way to
recover residual benzene. Some non-aromatic impurities will escape
with the vent gas, and most of it would be purged as liquid. This
methodology is described in U.S. patent application Ser. No.
08/957,252.
EXAMPLE
For illustration and consistency purposes, an ethylbenzene
production system for 380,000 tonne per year of ethylbenzene is
described. The streams and apparatus designations are depicted in
FIG. 1. The assumed production rate is based on 345 operating days
per year. The dilute ethylene feed (stream 1) to the facility is
originated from a naphtha based olefin plant. Stream No. 1, at a
pressure of 25 kg/cm.sup.2 -g and a temperature of 30.degree. C.,
has the following
Component kg-mol/hr Mol % Hydrogen 1,460 31.1 CO 21.0 0.44 Methane
2748.0 58.6 Ethylene 448.0 9.6 Ethane 9.0 0.2 Propylene 0.02 5 PPM
Acetylene 0.02 5 PPM Total 4686
Stream No. 2 contains impure benzene from a pyrolysis gasoline
source. More specifically, Stream No. 2 comprises:
Component kg/hr kg-mol/hr wt % Benzene 38,200 490 96.0 Cyclohexane
1,400 16.6 3.5 Dimethyl pentanes 160 1.6 0.4 N-heptane and C.sub.6
/C.sub.7 Trace Trace Trace Toluene Trace Trace Trace Water Trace
Trace 10 PPM Sulfur Compounds Trace Trace 0.5 PPM (as sulfur) Total
39,760 508
At the end of the run, there is a total ethylene utilization 98%
and 0.8% ethylene losses to heavy aromatics. Thus, 97.2% of the
ethylene is converted to ethylbenzene, and the balance is routed to
gaseous and liquid fuels. Impurities build up in the liquid of the
reflux drum 60 is 27 wt %. The system does not contain a purge
reactor. Ethylene enters the bottom of the ethylation catalyst bed
and reacts with benzene in liquid phase at 180.degree. C. The heat
of reaction, 12 MM Kcal/hr, about 975 Kcal per kg of ethylene is
mostly recovered by generating steam stream 16 and vaporizing
benzene. The benzene is recondensed at the overhead and the vent
gas chilling system 50. About 5.0 % of the ethylbenzene formed in
the ethylation catalyst beds further reacts with ethylene. About
3.8% of the ethylbenzene ends as DEB, 1.0% as TEB, and the balance,
0.2%, as heavier aromatics. The overhead of the ethylation catalyst
beds contains hydrogen, methane, benzene and small amounts of
ethylbenzene and unconverted ethylene. The overhead gas from the
ethylation beds proceeds to the rectifying section 10A, about 5 to
7 trays, where ethylbenzene is recondensed. Reactor vessel overhead
gas stream 3 proceeds to condenser/steam generator 20, and steam at
2.0 kg/cm.sup.2 -g is generated. The gas is further cooled in heat
exchange 30 to about 65.degree. C. by preheating tempered water
from 50.degree. C. about 80.degree. C. The gas is further cooled to
35.degree. C. in heat exchanger 40 with 30.degree. C. cooling
water. The vent gas at 35.degree. C. is chilled at heat exchanger
50 to -5.degree. C., by using +12.degree. and -8.degree. C.
refrigeration, for example, from the nearby olefin plant. Liquid
and vent gas products are separated in the reflux drum 60 and
reheated to 30.degree. C. by cold recovery at heat exchanger 50.
Vent gas, stream 13, at 22 kg/cm.sup.2 -g and 30.degree. C.
proceeds to PSA hydrogen recovery (not shown) or to a fuel gas
system at the following composition:
k-mol/hr Mol % Hydrogen 1,460 34.4 Methane 2,748 64.7 Ethane 9 0.2
CO 21 0.5 Ethylene 9 0.2 Benzene 4 0.1 Cyclohexane 1.2 --
Ethylbenzene Trace -- PEB Trace -- Total 4,252 100
A purge, stream 4, of 5,300kg/hr of liquid from the reflux stream
17 drum containing 73 wt % benzene is drawn from the reflux line at
30.degree. C. By applying the optional purge reactor (not shown)
the overall yield of ethylbenzene from benzene increases from 88%
to 97%. The benzene yield losses will show as pyrolysis gasoline if
no purge reactor is applied.
Liquid bottom product from the ethylation section 10-B, stream 12,
contains benzene; cyclohexane ethylbenzene and PEB descend to the
transalkylation section 10-C along with recycle DEB and TEB. About
80% of the PEB is reconverted to ethylbenzene at the start of the
run and 50% at the end of the run. The material balance is based on
end of run and transalkylation at 235.degree. C.
Transalkylated product line, internal stream 6, proceeds to the
benzene stripping section 10-D. Reboiler 80 provides the stripping
duty of about 16 MM Kcal/hr. Bottom product, stream 7, results from
benzene stripping and has the following composition:
kg/hr kg-mol/hr wt % Ethylbenzene 46,100 435 87.0 DEB 5,200 39 9.8
TEB 1,400 8.5 2.6 Heavy 130 0.6 0.25 Benzene 70 1.0 0.13
Cyclohexane 10 0.2 0.02 Total 52,910
The bottom stripped product, stream 7, at 310.degree. C., proceeds
to the ethylbenzene column 90. The overhead from the ethylbenzene
column, stream 8, is ethylbenzene product with 1,500 PPM of benzene
and 250 PPM of cyclohexane.
The bottom product from the ethylbenzene column stream 9
contains:
Ethylbenzene 200 kg/hr DEB 5,200 kg/hr TEB 1,400 kg/hr Heavy 130
kg/hr
This mixture, stream 9, proceeds to PEB column 100, where heavy
aromatics, stream 11, are separated as bottom product. DEB and TEB
overhead, stream 12, recycle to the transalkylation section 10-C
stream 14, by combining with benzene feed, stream 2.
In the conservative design, FIG. 2, the conversion in the
transalkylator will be 60% at liquid phase reactor at about
270.degree. C. About 3,500 kg/hr of DEB and TEB would react with
about 10,000 kg/hr of pure benzene feed from stream 15. The
material balance at FIG. 2 is somewhat different than FIG. 1 and
not shown.
* * * * *